Description of the Prior Art
Many liquid phase oxidation and hydrogenation reactions carried out in commercial operations are highly exothermic in nature. In such operations, the ability to remove the heat of reaction very often limits the production rate obtainable for a given reactor volume. Exothermic reaction with heat removal is typically accomplished in a stirred tank reactor with a cooling jacket, a stirred tank reactor with internal cooling coils, a stirred tank reactor with an external sidestream cooling system, or a bubble column reactor with heat transfer tubes. In all cases, the heat of reaction is transferred from hot reaction liquid through a solid surface into a cooler fluid such as cooling water, a refrigerant, or to evaporate water to make steam.
Heat transfer in all of these systems is described by the following equations EQU Q=rH.sub.r V (1); EQU Q=UA.DELTA.T (2);
and hence: EQU rH.sub.r =(A/V)U.DELTA.T (3);
where Q is the total heat load, r is the volumetric reaction rate, H.sub.r is the heat of reaction, V is the reactor volume, U is the overall heat transfer coefficient, A is the heat transfer surface area and .DELTA.T is the temperature difference between the reactor liquid and the heat transfer fluid. The left side of equation 3 is the volumetric heat of reaction or heat generation of the reactor, while the right side is the volumetric heat transfer capacity.
Equation 3 shows that heat generation increases with reaction rate and that, for steady state operation, the heat transfer capacity of the system must be increased when the reaction rate is increased. The equation also shows that heat transfer capacity is maximized when (1) the ratio of heat transfer area to reactor volume is maximized, (2) when the overall heat transfer coefficient U is maximized, and (3) when the temperature driving force .DELTA.T is maximized.
The area to volume ratio A/V is fixed by the geometry of the reactor and heat exchange system.
The heat transfer coefficient U is a function of fluid properties, and to a lesser extent the materials of construction of the heat exchanger. U can be increased or decreased by increasing or decreasing the flow rates of the reaction fluid and/or the heat transfer fluid. Cooling fluid flows are usually limited by pressure drop and in some cases by temperature considerations. Depending on the heat transfer system, the reaction liquid flow rate is limited by power input to the agitator or by pressure drop considerations.
The temperature difference .DELTA.T can be increased by increasing the reaction temperature and/or by decreasing the cooling fluid temperature. The reaction temperature is usually fixed so as to provide a given reaction rate and/or to minimize by-product formation. Thus, it is not usually desirable to raise the reaction temperature. The temperature of the cooling fluid is usually limited by the temperature of available cooling water, the cost of refrigeration or steam quality in evaporation systems.
In conventional reactor systems for exothermic reactions, a fundamental characteristic of jacketed reactor vessels is that they have a small A/V ratio. Since A increases as D, where D is the reactor diameter, and V increases as D.sup.2, A/V decreases as the size of the reactor is increased. Thus, jacketed reactors are typically used in small volume applications up to 100 gallons.
Stirred tank reactors with internal cooling coils typically have a higher ratio of A/V than jacketed vessels, particularly as the vessel sizes get larger. However, coils have several limitations. Heat transfer area is maximized by minimizing coil diameter, but pressure drop in the coil gives a lower limit for the coil diameter. It is possible to increase A/V by packing the reactor with coils. However, this tends to cause uneven flow distribution in the reactor which can lead to poor reactant mixing and undesirable by-product formation. It is also mechanically difficult to support multiple coils within the reactor vessel. Reactors with internal cooling coils are thus typically used in medium size applications of between 100 and 20,000 gallons. This reactor configuration is quite common in hydrogenation systems, such as in edible oil production, and in inorganic oxidations such as copper oxidation to copper sulfate.
One approach for precluding the geometric and flow constraints on the A/V ratio is to use a sidestream cooling system for external cooling. This kind of reactor configuration is often used in the oxidation of cumene to cumene hydroperoxide for the production of phenol, and in some hydrogenation systems.
In such sidestream systems, a sidestream from the reactor is pumped through a heat exchanger or other cooling system, and the cooled reaction liquor is returned to the reactor vessel. In principle, the ratio of A/V is not limited by constraints associated with reactor geometry. However, there are other potential problems associated with such systems. Since the cooling is accomplished outside the reactor vessel, the cooler will normally operate at a temperature that is significantly lower than the reaction temperature. Thus, proportionally more heat exchange area and/or coolant flow is required. Also, in gas-liquid reactor systems, such as for air oxidation, oxygen oxidation or hydrogenation reactions, the reactant gas must be prevented from entering the external cooling system. Gas tends to disengage from the liquid and collect in pockets in high spots in the exchanger and associated piping. This reduces the effectiveness of the heat exchanger. Gas can also collect in circulation pumps and cause cavitation or gas flooding of the pump. In oxidation systems where the gas is air or oxygen, any gas build up in the external heat exchange system could create an explosion hazard. It is possible to keep the gas out of the external heat exchange loop. However, in many cases, this causes increased by-product formation due to reactant starvation in reaction systems where good mixing of reactants is important.
The A/V of bubble column reactors is typically quite high. In one configuration, which is used in the production of organic acids, the bubble column is configured like a vertical shell and tube heat exchanger. The reaction takes place on the tube side while cooling fluid is circulated on the shell side. The gas is sparged into some of the tubes. Liquid circulation is caused by the gas lift effect of the gas bubbles in the sparged tubes. Thus, there is upflow and gas liquid contacting in those tubes that are sparged. There is downflow without gas liquid contacting in the remainder of the tubes. This configuration has two disadvantages. The liquid velocity on the reactant side of the tubes is limited to the bubble rise velocity which is normally between 1 and 5 ft/sec. This velocity limitation limits the heat transfer coefficient U. Also, the downflow tubes are not exposed to reactant gas. This may cause a lower volumetric reaction rate and/or by-product formation as a result of gas starved conditions in the down flow tubes.
Another common bubble column configuration, which is common in the so-called Witten process for producing dimethyl terephthalate by successive oxidation and esterification of p-xylene, is to use vertical tubes with reactant fluid on the outside and cooling fluid on the inside. This is mechanically very difficult to implement, but it is advantageous if the required reaction volume is large. In such systems, feed gas is sparged into the bottom of the reactor. The gas tends to collect into a plume such that there is a column of rising gas in one section of the reactor and ungased downflow in the remainder. This leads to the same conditions described above, namely limited heat transfer coefficient on the reactant side, and gas starved conditions in the downflow regions. In addition, since the reactant side flow is not uniform, this configuration may give rise to hot spots in the vicinity of the gas plume. These hot spots can cause undesired by-product formation due to over oxidation.
Mass transfer considerations are also very important, particularly in gas liquid reaction systems. If mass transfer is reaction rate limiting, reactor productivity is determined by it. Also, reactant starvation caused by mass transfer limitations can cause by-product formation which lowers chemical selectivity. It is well known that these problems occur in air based chemical oxidation systems.
Air based bubble columns and stirred tank reactor systems have inherent mass transfer limitations. Oxygen mass transfer is proportional to the oxygen concentration or oxygen partial pressure in the oxygen containing gas bubble. The concentration of oxygen in an air bubble in a bubble column or a stirred tank reactor is only 21% at the sparger. As oxygen dissolves into the reaction liquid, where it is consumed by the reaction, and as liquid evaporates into the air bubble, the oxygen partial pressure in the air bubbles decreases, while the partial pressure of nitrogen, which is a component of air, and the partial pressure of evaporated organic material increases. Thus, the mass transfer driving force associated with air is inherently lower than if pure oxygen is used as the reactant gas.
In conventional stirred tank and bubble column reactor designs, the oxygen partial pressure of the exiting waste air stream must be maintained below safety limits of 5% on an organic free basis in order to prevent formation of flammable gas mixtures in the reactor vapor space. Thus gas phase oxygen concentration in conventional reactor designs is constrained between 21% at the air feed point and 5% at the waste gas exit. In bubble columns, the air is injected at the bottom of the reactor. The gas bubbles rise through the liquid, and the gas phase oxygen concentration varies from 21% at the bottom to 5% at the top of the reactor. In a well mixed stirred tank reactor, the average oxygen concentration in the system is 5% throughout. Thus, for a given operating pressure, safety concerns in conventional reactor systems severely constrain the available mass transfer driving force. The situation can be improved somewhat by raising the overall system pressure which raises the oxygen partial pressure, or by purging the headspace with relatively large amounts of an inert gas such as nitrogen, but these alternatives are generally very expensive.
The net result of the limited mass transfer driving force inherent in conventional air based reactor systems is that oxygen starved conditions, and the accompanying product selectivity penalty, are more likely to occur as reaction temperature and reactor productivity are increased.
Another factor which limits oxygen mass transfer capacity is the degree to which the oxygen containing gas bubbles are uniformly distributed within the liquid phase. If some regions of the liquid phase are not exposed to oxygen containing gas bubbles, those regions will be oxygen starved and by-product formation will occur. Hence it is crucial to have good gas bubble distribution throughout the reactor.
In conventional bubble column reactors or gas lift bubble column reactors, the gas bubbles are introduced at the bottom of the reactor. They rise through the reaction liquid due to their buoyancy. The bubbles cause a recirculating liquid flow pattern. In bubble column reactors, the flow tends to be up through the center of the reactor and down near the walls of the reactor. The oxygen containing gas bubbles tend to concentrate in the center upflow region, which leaves the outer downflow region gas starved and subject to by-product formation reactions. In gas lift bubble columns the oxygen is typically sparged into gas lifted heat transfer tubes such that there is liquid upflow in the tubes. Additional tubes without spargers are provided for recirculating flow. Oxygen starved conditions and hence by-product formation reactions prevail in the downflow tubes.
One specific example of an oxidation system where heat transfer and mass transfer are critical is the production of aliphatic acids. Aliphatic acids are produced by liquid phase reaction of an aldehyde with oxygen according to the reaction: EQU R--HO+1/20.sub.2 =R--OOH
The aldehydes and corresponding acids, may be linear or branched, and the number of carbon atoms may vary from 3 to 12. The precursor aldehydes are often made using the Low Pressure Oxo (LPO) process. Hence, the derivative acids are often referred to as Oxo acids. The aldehydes may also be obtained or produced by means other than the LPO process, but this class of compounds is referred to as Oxo acids, none the less. The source of the aldehydes is not critical to this process.
In commercial production of such acids, selectivity to acid is typically between 80% and 99%. Selectivity decreases with chain length and the number of side chains or branches. For example, the selectivity of propionaldehyde which has three carbon atoms (C.sub.3), to propionic acid, is better than the selectivity of valeraldehyde to valeric acid, which has 5 carbon atoms (C.sub.5); and the selectivity of valeraldehyde, which is a linear five carbon molecule, to valeric acid is higher than the selectivity of 2-methyl butyraldehyde, which is a branched C.sub.5, to 2-methyl butyric acid. In commercial practice, by-product inhibitor additives may be added to some of these systems to improve selectivity.
In liquid phase aldehyde oxidation, the oxygen is typically introduced into the liquid by mass transfer from gaseous air bubbles. The oxidation reactions occur in the liquid phase; either in the bulk liquid phase or in the liquid film which surrounds the air bubbles. Oxygen starvation, that is lack of dissolved oxygen in the reaction liquid, promotes by-product formation reactions and hence reduces the selectivity of aldehyde to acid. Thus, adequate mass transfer of oxygen from the gas phase to the liquid phase is critical to maintain adequate dissolved oxygen concentration in the liquid phase in order to suppress by-product formation reactions.
It has been found that by-product formation increases with reaction temperature. Since reaction rate typically increases with temperature, the reaction consumes oxygen faster at higher temperature, and more oxygen is required to prevent the onset of oxygen starved conditions. Thus, gas-liquid mass transfer limitations become worse as temperature is increased, and therefore, it is more difficult to prevent oxygen starved conditions which cause by-product formation. The by-products formed under oxygen starved conditions are formate esters, ketones and alcohols.
Since the conversion of aldehyde to acid increases with temperature, it is possible to increase reactor productivity by increasing temperature. However, if the increase in temperature moves the reaction system into the oxygen starved regime, or makes an already oxygen starved condition worse, by-product formation reactions increase and selectivity to acid decreases.
Oxo acids are typically produced in air sparged stirred tank or bubble column gas lift reactors. At commercial reaction conditions, the exothermic heat of reaction produced by the oxidation reactions is significant. Although stirred tank reactors have been used for Oxo acid production, bubble column reactors configured as vertical shell and tube heat exchangers are preferred because of the higher A/V ratio.
In the bubble column reactors, air is sparged into the bottom of some of the heat transfer tubes, while the remainder of the tubes are not sparged. This combination of sparged and unsparged tubes causes a recirculating liquid flow within the reactor. The gas causes upflow of liquid in the sparged tubes, while downflow occurs in the remainder of the tubes that are not sparged. As air rises through the sparged tubes, oxygen transfers from the air into the liquid phase where it reacts with the aldehyde to form the acid. There is no mass transfer of oxygen into the liquid in the tubes which are not sparged.
In this reactor configuration, heat transfer occurs in all of the tubes and the ratio of A/V is high. However, the heat transfer coefficient U is limited somewhat because the tube side flow velocity is limited to the rise velocity of the gas bubbles which is typically between 1 and 5 ft/sec. Furthermore, since a fraction of the tubes are not sparged with gas these tubes operate in the mass transfer limited or oxygen starved mode. Thus, by-product formation is higher in the tubes which are not sparged compared to the tubes that are sparged. By-product formation is also favored by inherent mass transfer limitations associated with using air for the oxidant in the sparged tubes.
It will be appreciated from the above that improvements in the reactor system for oxidation, hydrogenation and other exothermic gas-liquid operations are highly desired in the art. Such improvements desirably would mitigate heat transfer limitations and improve mass transfer performance as compared to the conventional systems described above.
It is an object of the invention to provide an improved reaction system for oxidation, hydrogenation and other exothermic gas-liquid operations.
It is another object of the invention to provide a reactor system capable of mitigating heat transfer limitations and improving the mass transfer performance of exothermic gas-liquid operations.
With these and other objects in mind, the invention is hereinafter described in detail, the novel features thereof being particularly pointed out in the appended claims.